A Comparison Study between Batch and Continuous Process Simulation for the Separation of Carbon-13 Isotope by Cryogenic Distillation

Korean Chem. Eng. Res., Vol. 45, No. 1, February, 2007, pp. 57-66 Methanel C lk l i nj r o i p Çm Çmm Ç s 100-715 ne t 3 26 (2006 11o 15p r, 2006 12o 19p }ˆ) A Comparison Study between Batch and Continuous Process Simulation for the Separation of Carbon- Isotope by Cryogenic Distillation Jong Hwan Kim, Doug Hyung Lee, Euy Soo Lee and Sang Jin Park Department of Chemical Engineering, Dongguk University, 26, 3-ga Pil-dong, Chung-gu, Seoul 100-715, Korea (Received 15 November 2006; accepted 19 December 2006) k p rp }l d p 12 C l p ˆ oo C p k 1.1Í l p. Cp kr oo f p, k, p kl p v n tn o p. d p ˆ C p r p pp, n ˆ q lp p l p. krrp C oo rm v p rp v p p ~ w }l d oo p p CH 4 m H 2 Op pp l ll v CO v p. l l LNG NG C oo o l n r n l Rigorous rm v r Š m. m 12 CH 4 p p p n qk r Cp rp qlp r Feasibility Study n. SRK ˆ rep Acentric factor v k p l Acentric factor p l r Feasibility Studyl r r s p rm v Cp p p m pl. e l e rm v rp l v nr r p l rp rk m. l l Cp rm p rp nr p re pl. h Abstract Natural gases generally consist of mainly 12 C and about 1.1 of C. It is well known that a stable carbon isotope, C, has been widely used for the applications of medical, pharmaceutical, and agricultural tracers. As a result, the development of the separation and concentrating technology of C can cause of high value-added products and the possibility of the generation of new carbon materials. In general, there are two kinds of approaches to obtain a stable C isotope by the separation of cryogenic distillation. One is to obtain a concentrated isotope from natural gas. Another approach is to get concentrated CO by distillation followed by a chemical reaction of CH 4 and H 2 O. In this study, rigorous process simulations of the cryogenic distillation have been performed and analyzed for the concentrated separation of C isotopes from LNG and NG by using commercial process simulator. Due to the very small differences of relative volatilities and separabilities of 12 C and C, the process design and operation of effective separation and concentration of C need special strategies and feasibility studies. Utilization of vapor pressure data to acentric factor in SRK equation of state and optimized process conditions have been able to predict for the effective of the separation yield and concentration of C for the cryogenic distillation. The various operation strategies for both batch and continuous cryogenic distillation are also studied and suggested for the basic design of the process. Development of this study can provide a tool for the effective design and operation of the cryogenic separation of C. Key words: Carbon Isotopes, Cryogenic Distillation, Process Simulation, Carbon- To whom correspondence should be addressed. E-mail: sjpark@dongguk.edu 57

58 s Ëp Ëpp Ë v 1. v p ql l sq ˆ p 8 p ˆ oo ( 9 C~ C) l p p ˆ C(carbon-twelve) 16 12 sq p. p n ˆ oo p p p C (carbon-thirteen) p v ql l k 1.1Í r l pp lrp tn vp p. C 12 C m p kr oo,, kl, v, p, k,, l on n p (tracer, vp r j oo ) v n. n (urea), (glucose), (fructose), m p(triolein) p s t ˆ p C eˆ l s pn l k e~ svp p r v s l on pn p. p n, p m p p qp pp Œk p p qn p qp ~n CO 2 v l oo l p p r v p. v, v e p C Cp 12 r l svp ˆ d p p. m l, C urea pn l Žp (Helicobacter Pylori, oq r l p k) p p o ~ o kp opp p o kp kp okp v p p. CO 2, CH 3 OH, CF 3 Br p p ˆ oo C eˆ vp n d pr pn p. ql p ˆ l ˆ oo C l Cp p 12 p. m l, C 99Í eˆ rp C 99.99Íp p 12 v. C q r pn l n ep 12 p. p 12 C p pk rp lr p 1.1Íp C ql pk 50Í p p p. pm p q pn l p rq p l l m q(heat sink) n p, NMR l C 12 nr rn p. C p v n eqp l ˆ p q p. C p q q p l p p k p p ˆ l sq CO/ 12 CO oo rm v (cryogenic distillation) ~ p [1, 2]. p ˆ 99Í eˆ Cp p q e ˆ tl p. pnl k p ˆ(CH 4 )p rm v rp pn l oo rp n k r p. l p ˆ (CO 2 )l sq oo o e r l p ˆ d(ch 4 ) p oo l (thermal diffusion)p pn l p l pp, l ˆ d l rn p qn Ž qn p rp o l oo p p [3-5]. l p p d(carbon source) C l r q }l d rp. l p }l d LNG(liquified natural gas, t ˆ(CH 4 )) ˆ p l rn rn l n p. LNG o45 o1 2007 2k l o l p C / rp n n LNGp p p pp kn n q eqp } pp p. LNG C/ Cp oo rp l r 12 l n q Œq n. p pm p s n q lrp r nm p pv k p. p rp LNG NG oo p rmp s l f v p p. v k ˆp rm v ( v ) l p l rm l l p l p. v m 12 CH 4 p l p t (v k ) (separability, α)p p n qp rp ƒ r. p n rp p o bench scale p pilot scale e l pl q p (dimension) m m (m, wall effect, end effect, hydraulic effect ) e l ƒ m p. p e er m r o r(methodology) r p n, p m o p rp r p m kp re p [1-2, 6]. v l m p LNGp r p ˆ oo p rl t m p rp l l r q p ( e q )m p p o e p Š rp pl r l rp lk [7]. l l LNG / 12 CH 4 rm v r q o q m q p e l v l p / 12 CH 4 p e p q m. kn l l Cp r l r p v m pl rp rs p l rm v r l p Commercial Process Simulator pn l q m. 2. m n y 2-1. Methane lk m r h ˆp oo n o vp p rp t. l l Raoultp p /k p m. p o l v kp kkk l ˆ q pn m. ˆp v k p Antoineep p [6]. v ln P (P v L =.2093 897.84 L in kpa and T in K) (1) T ------------------ 7.16 p v kp pp ep. v log P L 36.9 0.192 ----- = --------- ------------ (T in K) (2) v T P H T 2 q op e (2)m p ˆ oo ~ p v kp 12 CH 4 ˆ p v k m, p p v kp p p p.

Methanep C oo o e l e rm v r l 59 Table 1. Comparison between α Eqn and α SRK for optimized acentric factor of α Eqn α SRK ω CH4 α SRK α Eqn α SRK -------------------------- 100 α Eqn 0.01112 1.003973 1.11 10 4 0.01105 0.011 1.004030 5.40 10 5 0.00538 0.01114 1.004088 4.00 10 6 0.00040 0.01115 1.004147 6.30 10 5 0.00627 0.01116 1.004204 1.20 10 4 0.01195 Fig. 1. Comparison of vapor pressure between 12 CH 4 & -I. Table 2. Relative volatility and K-value of 12 CH 4 / 12 CH 4 (Mole Fraction) K-value ( 12 CH 4 ) K-value ( ) Relative volatility 0.0 1.002940 1.000000 1.00294 0.2 1.002350 0.9994 1.00294 0.4 1.001760 0.998826 1.00294 0.6 1.001170 0.998240 1.00294 0.8 1.000590 0.997654 1.00294 1.0 1.000000 0.997069 1.00294 b = 0.08664 RT c -------- P c (6) Fig. 2. Comparison of vapor pressure between 12 CH 4 & - II. v P H 897.84.2093 T ------------------ = exp 10 7.16 0.192 36.9 ------------ --------- T T 2 (T in K) (3) e (1) e (3)l v v P L P H 12 CH 4 m p v kp ˆ. el k(101.325 kpa)l 12 CH 4 p rp 111.669800 Kp, p rp 111.704475 K p pp r p 0.034675 Kp. v, r p n p k p. op Fig. 1l 12 CH 4 m p m l v k p ˆ l 100 K p v k p Fig. 2l ˆ l. 2-2. SRK o l mk r h v k p r o SRK ˆ rep n l q m. SRK ˆ rep e (1)~ (4)m. RT at ( ) P = -------------- ------------------ ( v b) vv ( + b) 2 at ( ) 0.42747 R2 T = c -----------α ( T) P c (4) (5) α( T) 1 ( 0.48508 + 1.55171ω 0.156ω 2 0.5 = [ + )( 1 T r )] 2 op e l p b l p m m p k p p, T c, P c p p a(t) T c, P c m Reduced Temperature T r Acentric factorp p. l l p m l v kp r o 12 CH 4 m p T c, P c p p r, Acentric factor ω p Acentric factor rel v kp qkv Acentric factor v v ƒv 100 Kl e (2) α Eqn = P L P H = 1.004084p 100 Kl Commercial Process Simulatorl SRK v v ˆ rep pn l α SRK = K L K H m l m q q Acentric factor ω p m. p rp Table 1l ˆ l p s p Acentric factor ω p 0.01114 e. 12 CH 4 p Acentric factor ω p 0.01040 p. Table 2l 100 kpal 12 CH 4 p l 12 CH 4 m p K-value ˆ l. 100 kpal 12 CH 4 / p k 1.00294 pp k p. 2-3. Short-cut i m m l l p ˆ oo ~ p p 12 CH 4 m qnp p p lp ~p v qlp ˆ o pp p k~ (Ideal liquid mixture) r l r l. Shortcut p p n FUG(fenske-underwood-gilliland) p rn l v l n p m p q. 2-3-1. FUG(fenske-underwood-gilliland) FUG p pn p p v ˆp p pn l p p. p Korean Chem. Eng. Res., Vol. 45, No. 1, February, 2007 (7)

60 s Ëp Ëpp Ë v e (8)p Fenske ep pn l p. N min x D ( 1 x) ln -------------------------- D ( 1 x B ) x B lnsf = ------------------------------------ = ----------- lnα lnα (8) l l SF(separating factor) e (9)m p rp. x DLK SF ------------, xbhk, = ------------ x DHK, x BLK, (9) e (10)p Underwood ep pn l p. N α R j x Dj, min = ------------- 1 α j θ j = 1 (10) l l θ e (11) p. N α j x Fj, ------------ = 1 q α j θ j = 1 (11) e (8) e (9)l p m pn l Gilliand el p p p. Gilliand ep e (12). N N --------------- m 1 1 + 54.4ψ ---------------------------- ψ 1 = exp ----------- N+ 1 11 + 117.2ψ R R m ψ = --------------- R+ 1 ψ 0.5 (12) () e (8)~(11)p pn l / 12 CH 4 rm rl n p 1Íl 10Í rl N min = 747.4 R min = 140.37p 10Íl 80Í rl N min =967.3 R min = 87.16 p. 3. r o lp q lp p. ll p r p nn p tl s r p p o l n rp v rp. v r p v p p pn v p k~ p sqp o rp, rl p v pn p. sq k l kv, kv, rkv v v sq el e v, l e v. e v rp r nnl ol p v pl l e v r l s p r p v ˆl pp k s p o } p qrp v p. v l e v rl l v r p v p. Short-Cut Simulation Š SRK ˆ rep pn l Rigorous e v rp l p nr r p m. e v r p r Fig. 3l ˆ l. 3-1. r o Short-Cut Simulationp pn l }l d ˆ ˆˆp o45 o1 2007 2k Fig. 3. Schematic flow diagram for batch distillation. rm v rp }l d tl l p k 1.000Íp p 1.967Í v Batch Distillation Columnl eˆ rp m. r pn l r p 1,000 p tp, rp p m. p r l CH 4 Feed 1 kmol n r ˆrr p p 1.967 moleí m, ˆrr p p kp 0.5 kmol m. ˆ p k p 24 kpa, ˆrp k p 38 kpa l m. R/R min l, Heat Duty Table 3m. R/R min p 100r po Rp o l rp 100p p l Short-Cut Simulationp e m. e v r p o nr s p Table 4m. LNG t ˆp rn ~ ˆ p Table 3. Fenske/underwood calculation results of batch distillation R/R min N/N min Tarys (N) Reflux Duty (Mkcal/h) ratio Condenser Reboiler 50.500 1.010 1,003 12,000-12.58 12.58 75.250 1.007 1,000 18,000-18.75 18.75 100.000 1.005 998 24,000-24.91 24.91 124.750 1.004 997 30,000-31.08 31.08 149.500 1.003 996 35,000-37.25 37.25 Minimum reflux ratio (R min ) : 236.53, Minimum theoretical No. of trays (N min ) : 992.9 Table 4. Basic specification of batch distillation column Number of plate Top pressure Column pressure drop Vapor flow rate Feed Charge amount Composition 12 CH 4 Condenser temperature 500 24G[kPa] 14G[kPa] 20G[kgmol/h] 1G[kgmol] 0.99G[mol ] 0.01G[mol ] BubbleGTemperature

Methanep C oo o e l e rm v r l 61 Fig. 6. Batch distillation process operation strategy - Case 1. Fig. 4. composition vs. batch cycle time. Fig. 7. Batch distillation process operation strategy - Case 2. Fig. 5. composition vs. no. of stages. d. l p e l e v r l r} rp ˆ p ~ o n m. Feed Charge Table 4p nr s l Condenserp Holdupp Feed 50Í sr l nr e l Stillpotl Fig. 4l ˆ l. Fig. 4l nr e p 24e p p 1.76737Í p p p v k p k p. p e v r p nr e p 24e p r l m. p p Condenserp m -175.985 Cp o Heat Duty 0.0421 Mkcal/hp Stillpotp m -171.749 Cp o Heat Duty Condenserp Heat Dutym. Fig. 5l e v rp l r p. k 1,000 p p rlv p lt. p p p 1.95350 o Cp. 3-2. Condenser Holdupi r o p o e v r l p p Table 4p nr s l 1,000, nr e p 24e p s p l Condenserp Holdupp e r m. pl e rp r p Fig. 6 Fig. 7l ˆ l. Case 1(Fig. 6)l Condenserp Holdupp Feedp kp 50Í, Case 2(Fig. 7)l Condenserp Holdupp Feedp 82.217Í, Case 3, 4, 5l Condenserp Holdupp Feedp 65Í, 75Í, 80Í sr pr o veˆ o Distillate Flow Rate 0p p l r l Stillpotl k kk k. Fig. 8l Case 1l p nr e l tn p ˆ l. Fig. 8. composition of each stage vs. batch operation time. Korean Chem. Eng. Res., Vol. 45, No. 1, February, 2007

62 s Ëp Ëpp Ë v Table 5. Condenser : stillpot = 50 : 50 (binary system) - Case 1 Feed Stage 1 Stage 2 Stage 3 Stage 4 Product yield ( ) 100.000 50.000 25.000 12.500 6.250 purity ( ) 1.00000 1.95350 3.81457 7.44221 14.49248 Table 6. Condenser : stillpot = 82.217:17.783 (binary system) - Case 2 Feed Stage 1 Stage 2 Stage 3 Stage 4 Product yield ( ) 100.000 17.783 3.149 0.556 0.093 purity ( ) 1.00000 4.93071 23.84666 85.40621 99.43025 Table 5l p Case 1(Fig. 6)p l p 1 Í r l ee r 4 14.49248Í Product p 6.25Í. p p Condenser p m -175.985, Reboiler(Stillpot)p m -171.744 l. e v rp nrp Fig. 6 p 4 (stage)p rp r p. 1 (8 ), 2 (4 ), 3 (2 ), 4 (1 ) rp 15, v 15dayp nr e p 1.000Íp Feed s p v l 14.49248Íp 0.5 kmolp s Product lp p. Table 6l p Case 2(Fig. 7)l p p 1.000Í r l r 4 99.43025Í Product p Feedp k 0.093Í. Fig. 7 p 192 +34 +6 +1 =233 p, v 233dayp nr e p 0.17783 kmolp lp p p p 99.43025Í. v Case 3, 4, 5l Condenser Holdupp Case 1 Case 2p t vrp 65Í, 75Í, 85Í l Feedp p 1.000Í r l r 4 51.10115Í, 95.38173Í, 98.86062Í Product p 1.346Í, 0.391Í, 0.160Í. Case 3, 4, 5p l p Product Yield p Table 7~9l ˆ l r p Fig. 9~11l ˆ l. l v Condenser Holdupl r nkp Table 10 Fig. 9. Batch distillation process operation strategy - Case 3. Fig. 10. Batch distillation process operation strategy - Case 4. Table 7. Condenser : stillpot = 65:35 (binary system) - Case 3 Feed Stage 1 Stage 2 Stage 3 Stage 4 Product yield ( ) 100.000 35.000 12.115 4.038 1.346 purity ( ) 1.00000 2.70586 7.30668 19.60361 51.10115 Table 8. Condenser : stillpot = 75:25 (binary system) - Case 4 Feed Stage 1 Stage 2 Stage 3 Stage 4 Product yield ( ) 100 25.000 6.250 1.563 0.391 purity ( ) 1.00000 3.65380.26314 46.40834 95.38173 Table 9. Condenser : stillpot = 80:20 (binary system) - Case 5 Feed Stage 1 Stage 2 Stage 3 Stage 4 Product yield ( ) 100 20.000 4.000 0.800 0.160 purity ( ) 1.00000 4.44875 19.52951 74.29939 98.86062 o45 o1 2007 2k Fig. 11. Batch distillation process operation strategy - Case 5. l ˆ l Product Yield, r, nr e p l Condenser Holdupp rr sr l nr lk p. v p n n Condenser Holdupp r k rp nr e p n, r p pk Product n n Condenser Holdupp q k p. 4. i r o Short-Cut Simulation Š SRK ˆ rep pn l Rigorous l e v rp

Methanep C oo o e l e rm v r l 63 Table 10. Results of batch process simulation for various condenser holdup Condenser holdup ( ) TotalGfeed (kmol) Final product (kmol) comp. of final product (mol ) Product yield ( ) (Feed/Product) Total No. of batch operation (Operation time, day) 50.000 8 0.500 14.492 6.250 8 + 4 + 2 + 1 = 15 65.000 26 0.350 51.101 1.346 26 + 9 + 3 + 1 = 39 75.000 64 0.250 95.382 0.391 64 + 16 + 4 + 1 = 85 80.000 125 0.200 98.861 0.160 125 + 25 + 5 + 1 = 156 82.217 192 0.178 99.430 0.093 192 + 34 + 6 + 1 = 233 nr r m p r p p kk k. 4-1. i r o Short-Cut Simulationp pn l }l d ˆ ˆˆp rm v rp }l d tl l p k 1.000Íp p 10.000Í v Column 1l eˆ rp m. r pn l r p 1,000 p tp, rp p m. p r l CH 4 Feed 1 kmol/h n r ˆ r p p 0.500 moleí m, ˆrr p p 10.000Í m. ˆ p k p 24 kpa, ˆrp k p 38 kpa l m. R/R min l, Heat Duty Table 11. Column 2 l k 10.000Íp p 80.000Í v eˆ rp m. r pn l r p 1,000 p tp, rp p r m. p r l 10Í Feed 0.05 kmol/h n r ˆ r p p 6.800 moleí m, ˆrr p p 80.000Í m. ˆ p k p 24 kpa, ˆrp k p 38 kpa l m. R/R min l, Heat Duty Table 11. Fenske/underwood calculation results of column 1 Tarys Reflux Duty (Mkcal/h) R/R min N/N min (N) Ratio Condenser Reboiler 1.500 1.547 1,156 210.555-0.4222 0.4222 1.750 1.427 1,066 245.648-0.4922 0.4923 2.000 1.348 1,008 280.740-0.5623 0.5623 2.250 1.293 966 315.833-0.6323 0.6323 2.500 1.252 936 350.925-0.7023 0.7024 Minimum reflux ratio (R min ) : 140.37, Minimum theoretical No. of trays (N min ) : 747.4 Table 12m. Table 11 Table 12l lt p Column 1l k 1Íp p 10Í rp R/R min 2 r N/N min k 1.348 nr 1,000, Reflux Ratio 300p Column 2l k 10Íp p 80Í rp R/R min 15 r N/N min k 1,035 nr 1,000, Reflux Ratio 1,500p m. Column 2p Reflux Ratio 1,500 p po Feed Rate o qp kp Columnp Sizingl r (80Í p )p o p. 4-2. i r o ~ w l e v r o nr s p Table Table. Specification of continuous distillation column Column 1 Number of trays Feed tray Top pressure Prssure drop Feed Rate Composition 12 CH 4 Condenser Reflux ratio [Mol Basis] Column 2 Number of trays Feed tray Top pressure Prssure drop Condenser Reflux ratio [Mol Basis] 1,000 50 24G[kPa] 14G[kPa] 1G[kgmol/h] 99.00G[mol ] 1.00G[mol ] 300 1,000 50 24G[kPa] 14G[kPa] 1,500 Table 12. Fenske/underwood calculation results of column 2 R/R min N/N min Tarys (N) Reflux Duty (Mkcal/h) Ratio Condenser Reboiler 8.000 1.067 1,033 697.291-0.0703 0.0703 11.500 1.046 1,012 1,002.356-0.1011 0.1011 15.000 1.035 1,001 1,307.421-0.18 0.18 18.500 1.028 994 1,612.486-0.1625 0.1625 22.000 1.024 990 1,917.550-0.1933 0.1933 Minimum reflux ratio (R min ) : 87.16, Minimum theoretical No. of trays (N min ) : 967.3 Fig. 12. Schematic flow diagram for continuous distillation. Korean Chem. Eng. Res., Vol. 45, No. 1, February, 2007

64 s Ëp Ëpp Ë v Table 14. Feeds and products of column - Case 1 Stream name FEED OVHD1 BTMS1 OVHD2 BTMS2 Phase Liquid Liquid Liquid Liquid Liquid Temperature ( C) -180.000-175.985-171.746-175.983-171.719 Pressure (kpa) 52.000 24.000 38.000 24.000 38.000 Flow rate (kmol/h) 1.0000 0.95000 0.05000 0.04750 0.00250 Molecular weight 16.053 16.048 16.147 16.110 16.841 Composition CH 4 0.9900000 0.9950165 0.8946717 0.9317066 0.1919253 0.0100000 0.0049835 0.1053283 0.0682934 0.8080747 Table 15. Heat duty and temperature of condenser and reboiler and column diameter - Case 1 Column name Column 1 Column 2 Condenser duty (Mkcal/h) -0.6040-0.1500 Reboiler duty (Mkcal/h) 0.6041 0.1503 Condenser temperature ( C) -175.985-175.983 Reboiler temperature ( C) -171.746-171.719 Diameter (mm) 1,219 628 r Fig. 12l ˆ l. p Columnp Bottom Product Rate Columnl Total Feed Ratep 5Í m, Rebolierm Condenserp Heat Duty l Reflux Ratio 300, 1,500p seˆ ˆ p k p 24 kpa, ˆrp k p 38 kpa p l r ee m. Column Diameter Structured Packingp Sulzerp Mellapak 125.Y p m. ~ w r Table 14m Table 15l ˆ l. s Product p k 80.81Í 0.00250 kmol/hp Rate p k pl. p 1 nr r n 300 p 24 h = 7,200 hp 7,200 h 0.00250 kmol/h 16.841 = 303.1 kg p 80.81 molí p p 303.1 kg/yr p p k pl. Columnp Diameter 1,219 mmm 628 mm l. w l e v r o nr s p Table r Fig. l ˆ l. p Columnp Bottom Product Rate Columnl ~ w np Flow ratem p Rate v, 0.0500 kmol/h, 0.0025 kmol/h r m, ~ w np OVHD2 Streamp s p p k 6.829Íp p p. Feedp p 1.000Í p Fig.. Schematic flow diagram for continuous distillation with recycle stream. Recycle eˆ Rebolierm Condenserp Heat Duty l Reflux Ratio 300, 1,500p seˆ ˆ p k p 24 kpa, ˆrp k p 38 kpa p l r ee m. Column Diameter ~ w nm v Structured Packing p Sulzerp Mellapak 125.Y p m. Table 17. Heat duty and temperature of condenser and reboiler and column diameter - Case 2 Column name Column 1 Column 2 Condenser duty (Mkcal/h) -0.6325-0.1502 Reboiler duty (Mkcal/h) 0.6325 0.1502 Condenser temperature ( C) -175.985-175.981 Reboiler temperature ( C) -171.741-171.7 Diameter (mm) 1,248 628 Table 16. Feeds and products of column - Case 2 Stream name FEED OVHD1 BTMS1 OVHD2 BTMS2 Phase Liquid Liquid Liquid Liquid Liquid Temperature ( C) -180.000-175.985-171.741-175.981-171.7 Pressure (kpa) 52.000 24.000 38.000 24.000 38.000 Flow rate (kmol/h) 1.0000 0.99750 0.05000 0.04750 0.00250 Molecular weight 16.053 16.051 16.204 16.166 16.925 Composition CH 4 0.9900000 0.9922489 0.8364832 0.8749315 0.1059294 0.0100000 0.0077511 0.1635168 0.1250685 0.8940706 o45 o1 2007 2k

Methanep C oo o e l e rm v r l 65 Table 18. Feeds and products of column - Case 3 Stream name FEED OVHD1 BTMS1 OVHD2 BTMS2 Phase Liquid Liquid Liquid Liquid Liquid Temperature ( C) -180.000-175.985-171.743-175.983-171.716 Pressure (kpa) 52.000 24.000 38.000 24.000 38.000 Flow rate (kmol/h) 1.00000 0.99564 0.06807 0.06372 0.00435 Molecular weight 16.053 16.049 16.157 16.111 16.838 Composition CH 4 0.9900000 0.9935156 0.8844319 0.9314984 0.1955796 0.0100000 0.0064845 0.1155681 0.0685016 0.8044204 Table 19. Heat duty and temperature of condenser and reboiler and column diameter - Case 3 Column name Column 1 Column 2 Condenser duty (Mkcal/h) -0.63-0.2015 Reboiler duty (Mkcal/h) 0.6314 0.2014 Condenser temperature ( C) -175.985-175.983 Reboiler temperature ( C) -171.743-171.716 Diameter (mm) 1,246 720 w r Table 16m Table 17l ˆ l. s Product p k 89.41Í 0.00250 kmol/hp Rate p k pl. p 1 nr r n 300 p 24 h = 7,200 hp 7,200 h 0.00250 kmol/h 16.925 = 304.7 kgp 89.41 molí p p 304.7 kg/yr p p k pl. Columnp Diameter 1,248 mmm 628 mm l. Recycle Flow sq Column 1p Diameter ~ w n k ƒv p k pl. w l e v r o nr s p w s r Fig. l ˆ l. p Columnp Bottom Product Rate Columnl ~ w np Productp p m d Flow rate 0.0681 kmol/h, 0.00435 kmol/h m, OVHD2 Streamp Recycle eˆ Rebolierm Condenserp Heat Duty l Reflux Ratio 300, 1,500 p seˆ ˆ p k p 24 kpa, ˆrp k p 38 kpa p l r ee m. Column Diameter ~ w nm v Structured Packingp Sulzerp Mellapak 125.Y p m. w r Table 18 Table 19l ˆ l. s Product p k 80.44Í 0.00435 kmol/hp Rate p k pl. p 1 nr r n 300p 24h=7,200 hp 7,200 h 0.00435 kmol/h 16.838 = 527.4 kg p 80.44 molí p p 527.4 kg/yr p p k pl. Columnp Diameter 1,246 mmm 720 mm l. Bottom Rate w n l Column 2p Diameter w n s ƒv p k pl. ~ w n w, w nl Recycle Streamp sq l Column Diameter ƒv Duty v v w n p Product Flow Ratep ~ w n p Product lp p w nl ~ w nm d p v p v p k p. p r Table 20. 5. l l LNG / 12 CH 4 rm v r q o q m q p e l v l p / 12 CH 4 p e p q m. SRK ˆ rep Acentric factor l v p m p rn l e v r l r r r o Condenserp Holdupl r m. Condenser Holdupp k 82Í p p n 99Í p p p p s Product lp pl. l e v r Columnl Bottoml s k pl k 1,000 p 2 Columnp l e v rp p Product Rate rp Recycle o l k 80Íp Table 20. Results of continuos process simulation for various case Case 1 Case 2 Case 3 ColumnG1 ColumnG2 ColumnG1 ColumnG2 ColumnG1 ColumnG2 Condenser duty (Mkcal/h) -0.6040-0.1500-0.6325-0.1502-0.63-0.2015 Condenser temp. ( C) -175.985-175.983-175.985-175.981-175.985-175.983 Reboiler duty (Mkcal/h) 0.6041 0.1503 0.6325-0.1502 0.6314 0.2015 Reboiler temp. ( C) -171.746-171.719-171.741-171.7-171.743-171.716 Column diameter (mm) 1,219 628 1,248 628 1,246 720 Final product composition ( ) 80.807 89.407 80.442 Final product rate (kmol/h) 0.00250 0.00250 0.00435 Product amount per year (Kg/yr) 303.1 304.7 527.4 Korean Chem. Eng. Res., Vol. 45, No. 1, February, 2007

66 s Ëp Ëpp Ë v l Recyclep lp n 303.1 kg/h, Recyclep pp n 527.4 kg/h p k pl. Recyclep p rl Recyclep l r p p r l n p k 89Í p k pl. CO/ 12 CO p rm v rl r l LNG pn ˆ oo rl mp e l r e np p re pl. y 1. Johns, T. F., Vapor Pressure Ratio of 12 C 16 O and C 16 O, Proc. Phys. Soc., B66, 808-809(1953). 2. McInteer, B. B., Isotope Separation by Distillation: Design of a Carbon- Plant, Sep. Sci. Tech., 15(3), 491-508(1980). 3. A Method for the Separation of Carbon Isotopes by Chemical Exchange Method, Europ. Pat. 0042877(1980). 4. Vasaru, G., Ghete, P., Cocacci, I. and Atanasiu, M., Separation of Carbon- by Thermal Diffusion, Stable Isotope Life Sci. Proc. Tech. Comm. Meet Mod Trends Biol. Appl. Stable Isot., 39-52(1977). 5. Method of Centrifugal Enrichment of Carbon- Isotope in CO2 Form, RU 2153388(2000). 6. Hirose, Y., Tachibana, H. and Soh, H., Distillation Calculation Method for Isotope Separation, Kagaku Kogaku Ronbunshu, 223(3), 527-533(1996). 7. Douglas, J. M., Conceptual Design of Chemical Processes, McGraw-Hill Book Co.(1988). 8. Xopowulov, A. B., Separation of Stable Carbon Isotopes by Cryogenic Distillation, Khimicheskay Promyshlennost Moscow, 4, 229-335(1999). 9. Smith, J. M., Van Ness, H. C. and Abbott, M. M., Introduction to Chemical Engineering Thermodynamics, McGraw-Hill Book Co.(1996). 10. Prausnitz, John, M., Lichtenthaler, Rüdiger N., Edmundo Gomes de Azevedo, Molecular Thermodynamics of Fluid-Phase Equilibria, Prentice Hall(1999). 11. Tokyo Gas Report, EIT No. 32, The Utilization of LNG Cryogenic Energy(2000). 12. Alberto Bertuccon and Cristina Mio, Prediction of Vapor-liquid Equilibrium for Polymer Solution by a Group-contribution Redlich- Kwong-Soave Equation of State, Fluid Phase Equilibria, 117(1-2), 18-25(1996).. King, C. J., Separation Processes, McGraw-Hill Book. Co.(1980). 14. José O. Valderrama and Alexis Silva, Modified Soave-Redlich- Kwong Equations of State Applied to Mixtures Containing Supercritical Carbon Dioxide, Kor. J. Chem. Eng., 20(4), 709-715(2003). 15. Jungho Cho, A Study on the Simulation of Toluene Recovery Process using Sulflane as a Solvent, Kor. Chem. Eng. Res., V44(2), 129-5(2006). o45 o1 2007 2k